Starting up procedure in the hydrocaracking of hydrocarbons



Oct. 19, 1965 KLlNE ETAL 3,213,012

STARTING UP PROCEDURE IN THE HYDROCRACKING OF HYDROCARBONS Filed Sept. 25, 1962 640 E 3 620 QQ (k 600 2 4 a axe/2 l4/6 lazo DAVS 0F OPERAT/OA/ Fig. .1

k Mk fig, 680 0Q 13% R lu 640 504 1/2 A QM 35% 600 :k 15 /7 /9 2/ 2a 27 29 3/ .aa 04y: 0F OPEQAT/ON Pig. 2

u 720 Q0 N0 50mm 1100/ 680 w m SULFU gig 640 R DAYS OF OPERATION INVENTORS F 1, 5 BY 2059?? E. KLl/VE Wk ATTORNEY United States Patent STARTING UP PROCEDURE IN THE HYDRO- CRACKlNG 0F HYDROCARBONS Robert E. Kline, Pittsburgh, and Joseph B. McKinley,

New Kensington, Pa., assignors to Gulf Research 8: Development Company, Pittsburgh, Pa., :1 corporation of Delaware Filed Sept. 25, 1962, Ser. No. 226,090

16 Claims. (Cl. 208110) This invention has for its object to provide improved procedure for the hydrocracking of distillate hydrocarbons.

Hydrocracking of distillate hydrocarbons in certain respects is a relatively expensive operation. In order to improve the economy of the process, one of the main requirements is to use a highly active catalyst which will retain high activity over a long period of time. While highly active catalysts have been employed, there is still room for improvement in this regard. Furthermore it is frequently found that the more active the catalyst the shorter the onstream period. This is not due to any inherent defect in the catalyst but to improper procedure in employing the catalyst or in controlling the reaction conditions.

This invention has for its object to provide improved hydrocracking procedure. Another object is to provide improved hydrocracking procedure wherein a catalyst containing nickel-tungsten is employed. A still further object is to provide an improved procedure for hydrocracking distillate stocks whereby the hydrocracking process can be carried out over longer periods of time at high conversion rates. Another object is to provide improved start-up procedure. Other objects will appear hereinafter.

These and other objects are accomplished by our invention which includes subjecting a distillate feed stock composed primarily of hydrocarbons boiling in the range be tween about 350 and 800 F. to treatment with hydrogen in the presence of a nickel-tungsten sulfide catalyst composited with a siliceous carrier having a cracking activity index of at least 40. (See J. Alexander et a1. Laboratory Method for Determining the Activity of Cracking Catalysts, National Petroleum News, vol. 36, 1944, p. R537.) The catalyst and feed stock are initially contacted at below 500 F. and while the nickel and tungsten are substantially unsulfided, i.e. in the reduced, partially reduced or preferably in the oxide form. Thereafter the temperature is increased at a rate which avoids localized zones of excessive temperature with resultant localized deposition of a relatively high percentage of coke. This operation is continued until the desired conversion is reached. Thereafter the temperature is increased at a rate sufficient to maintain conversion at the desired level. Sulfur or a sulfur compound is introduced so that the catalyst will become substantially sulfided after initial contact of feed with the catalyst, but prior to expiration of the major part of the conversion period. We have found in accordance with our invention that a catalyst which is presulfided and maintained in that form by sulfur addition to the feed has less activity than a catalyst which is substantially in the oxide and/or reduced form and is started up as specified.

The catalyst utilized in our process may contain between about and 40 percent and preferably 10 to 25 .percent of nickel plus tungsten (determined as metals). The

3,213,012 Patented Oct. 19, 1965 "ice atomic ratio may be between 1 atom of tungsten to 0.1 atom of nickel to 1 atom of tungsten to 5 atoms of nickel. We prefer a range of between 1 atom of tungsten to 0.3 atom to 4 atoms of nickel. Any siliceous carrier may be employed which has .an activity index of at least 40. We preferably employ a siliceous carrier which has an activity index above 45. Such siliceous carriers are known in the catalytic cracking art, a typical example being silica-alumina cracking catalysts. Such catalysts may contain varying amount of alumina and silica and also may contain other metal oxide components such as zirconia, thoria, magnesia, etc. Any of such mixtures may be employed as long as the carrier displays the characteristics of an active silica-alumina cracking catalyst. Especially useful materials are the so-called high alumina synthetic silica-alumina cracking catalysts containing 25 percent alumina such as Triple A silica-alumina or the equivalent catalyst in pelleted form. These carriers are most advantageous when they are characterized by a surface area of 450 square meters per gram or greater, and preferably 500 square meters per gram or greater.

The catalyst employed in our invention may be prepared using any known procedure for manufacture of such multi-component catalysts. Thus the nickel and tungsten components may be deposited upon the cracking carrier by co-precipitation. Alternatively they may be deposited in sequence with or without intervening calcining. Simultaneous impregnation from a two-component solution containing the two metals may also be employed. Thus the procedure described in McKinley et al. Patent 2,703,789 would be entirely satisfactory. These components may be present as mixtures and/or as chemical compounds.

It is necessary that at the time the feed stock is first contacted therewith the nickel and tungsten components of the catalyst employed in our invention be largely unsulfided, i.e. largely in the reduced, partially reduced and/or oxide form. It is preferable to start out with a catalyst which is substantially in the oxide form. If a presulfided catalyst is used initially and kept in that form by sulfur addition to the feed, we have found that the useful life of the catalyst is substantially shorter than if a substantially unsulfied catalyst is initially employed. However, it is to be understood that a presulfided and subsequently reduced or largely reduced catalyst is satisfactory to use in the process of our invention.

Sulfiding of the catalyst may be accomplished by adding elemental sulfur or suitable sulfur compounds to the fresh feed, recycle feed, make-up hydrogen and/or the recycle hydrogen rich gas stream. The sulfur added includes the relatively small amount usually present in normally hydrogen refined or other feed and the relatively large amount which builds up in any hydrogen rich recycle gas stream when employing sulfur addition. Normally there is no sulfur in the liquid recycle but if any is present it is also included. Of course the operation may be single pass, in which case only sulfur added to and in the fresh feed and the process gas comes into consideration. Any organic or inorganic sulfur compound having a hydrogen-to-sulfur or a carbon-to-sulfur linkage as well as elemental sulfur can be used such as butyl mercaptan, thiophene, hydrogen sulfide, carbon disulfide, etc. It is convenient and often preferable to obtain rapid initial sulfiding and therefore it is advantageous to use large amounts of sulfur (to be hereinafter understood as designating elemental sulfur or any sulfur compound having carbonor hydrogen-to-sulfur linkage) during the early low temperature stages and then reduce the amount of sulfur to that which will maintain the catalyst in substantially sulfided condition. Thus it is advantageous to add 0.1 to 2.0 percent sulfur (determined as elemental sulfur when a sulfur compound is used) based on total liquid hydrocarbon feed (fresh feed plus any recycle) during early stages of the onstream cycle and reduce the sulfur to between about 40 p.p.m. and 1.5 percent later on and especially during the 55 to 80 percent conversion period. (Conversion means the actual percent conversion of feed to material boiling below the initial boiling point of the feed, which term includes unconverted feed recycle in recycle operation.) While we prefer to utilize a high sulfur content during early stages of the onstream cycle in order to cause rapid sulfiding and conditioning of the catalyst, we can utilize any sulfur content in feed, recycle hydrogen, make-up hydrogen and/or recycle feed which will result in substantial sulfiding and conditioning of the catalyst after intial contact of feed and catalyst but prior to expiration of the major portion of the conversion period. This of course includes simultaneous initial feed contact and sulfiding. Sulfur contents in the feed, hydrogen, etc. based on total liquid hydrocarbon feed may vary from about 40 p.p.m. to 2.0 percent. Larger amounts of sulfur than 2.0 percent can of course be used but there does not appear to be any economic advantage in using them. Straight-fun feeds usually require less than about 1000 p.p.m. of sulfur during the conversion period. Aromatic feeds require larger amounts of sulfur than straight-run feeds. Sulfur addition allows a given nitrogen content feed to be processed at a lower pressure or a higher nitrogen content feed stock to be processed at a given pressure with good catalyst aging. In contrast to the amount of sulfur we use, normally hydrogen refined feeds contain only a small amount if any sulfur.

Complete theoretical catalyst sulfiding usually does not take place. For instance, even with the addition of large amounts of excess sulfur, the catalyst in many cases is sulfided only to about 65 percent of theoretical. Such operations are understood to be within the scope of our invention. Although we refer to sulfides, sulfiding, etc, this is not to be taken to necessarily indicate the chemical form in which the hydrogenating components are present. Thus in accordance with our invention the hydrogenating components may be present as mixtures of the sulfides and/ or in the form of chemical combinations such as nickel thio tungstate.

The preferred modification of our process uses a catalyst combined with a halogen, preferably fluorine. Advantageously this halogen is incorporated in the catalyst during preparation by means of a compound such as HF; NH F; NH F-HF, H SiF or HBF or corresponding or similar compounds of chlorine or bromine such as hydrochloric acid, etc. Alternatively it is possible to halogenate the catalyst by adding a halogen compound such as difluoroethane, ortho-fiuorotoluene, fluorine, hydrogen fluoride, a sulfur fluoride, etc. to the liquid or gaseous feed but this method is not necessarily equivalent due to initial use of the catalyst in its unhalogenated and nonoptimum state. Most advantageously this latter method of adding halogen is employed for maintaining the halogen content of catalysts or for making minor adjustments in the catalysts halogen content. About 0.1 to 5 percent halogen may be present in the catalyst as the result of prior addition or addition during the onstream period. We prefer to use a catalyst containing between 0.5 and 3.5 percent halogen and especially between 1.5 and 2.5. When adding halogen to the catalyst during the onstream period it has been found desirable for economic reasons to limit the addition rate to an amount less than 20 p.p.m. halogen or halogen compound calculated as elemental halogen based on total liquid reactor feed and preferably to less than 10 p.p.m. When adding halogen to maintain the catalysts halogen content, it is advantageous to practice continuous addition. Usually 3 to 4 p.p.m. halogen in the total liquid reactor feed will accomplish the desired catalyst halogen level retention. If discontinuous halogen addition is practiced it is desirable to make the addition periods sufficiently frequent that in excess of 10 p.p.m. halogen, based on total liquid reactor feed, does not have to be employed. Nonmaintenance of the halogen content of the catalyst leads to a less selective reaction and contributes to shorter cycle operation.

Our invention is applicable to the hydrocracking of any distillate feed which is composed predominantly of liquid hydrocarbons boiling below 800 F. and generally between about 350 F. and 800 F. at reaction temperatures ranging generally from about 400 to 800 F. but more suitably from about 450 to 750 F. The feed stocks will be substantially free of asphaltic materials in view of the fact that they are obtained by distillation. Subsequent contamination of feeds with asphaltic material is undesirable. The feed stock may be of straightrun or cracked origin. These feeds should contain no more than 25 p.p.m. nitrogenous materials determined as nitrogen and preferably between 0 and 5 p.p.m. to be useful for the lower pressure aspects of this invention, i.e. at hydrogen partial pressures below about 1500 p.s.i. If they contain in excess of this amount they should be pretreated in some suitable fashion as for example hydrogenation to effect nitrogen reduction.

It is preferable for economic reasons to use as low a pressure as possible consistent with long cycle length. Usually about a three-months cycle length before initial catalyst regeneration is about the minimum which will provide an economic process so far as ultimate catalyst life and other process features are concerned. When operating according to the process of this invention this is easily obtainable with all feeds whether of cracked or straight-run origin at the relatively high hydrogen partial pressures of 1500 to 3000 p.s.i. even though all process conditions are not optimum. If, however, process conditions and procedures are as described herein, this minimum desirable cycle length can be obtained at hydrogen partial pressures of between about 400 to 1500 p.s.i. Thus a hydrogen partial pressure of about 1000 p.s.i. is adequate for the processing of a cracked stock such as a 450 to 650 F. boiling range hydrogen refined catalytically cracked furnace oil distillate, whereas, a hydrogen partial pressure in the range 400 to 600 p.s.i. is adequate when treating a hydrogen refined 350 to 550 F. light straight-run furnace oil distillate.

It is important that the initial temperature when the feed stock and hydrogen first contact the catalyst be below about 500 F.; otherwise due to the high activity of the fresh or regenerated catalyst, local zones of excessive temperature will rapidly deactivate or otherwise cause damage to the catalyst. It is preferable, although not necessary, to employ an initial temperature which is above the dew point of the feed under conditions existing in the reactor. After initial contact has taken place, the temperature may be increased until 55 to percent conversion is reached. However, the rate of increase should be such that no local zones of excess temperature are formed with simultaneous coking of the feed and concomitant lowering of activity of the catalyst. Ordinarily this breaking-in period will require only about one or two days. While this cannot be specified in precise arithmetic terms since it will depend upon the individual catalyst, the specific reaction conditions and the feed stock, in general, a temperature increase of 10 F. per hour has been found satisfactory during this breaking-in period. The rate of temperature increase during this initial period will be substantially greater than the rate of increase which is employed after the desired conversion is reached. Thus the 55 to 80 percent conversion period usually will last for several months. After the reactor has been brought to a temperature at which 55 to 80 percent conversion is obtained, the catalyst is then increased in temperature but only at a rate suificient to maintain this degree of conversion. Since the reaction is exothermic, the temperature within the reaction zone may be controlled by lowering or raising the inlet temperature of the feed and/or hydrogen and/or by injecting cooling hydrogen or liquid feed into the reactor.

When a point is reached where an excessive rate of temperature increase is necessary to maintain desired conversion such as at 55 to 80 percent or when a maximum temperature of about 800 F. or more generally about 750 F. is reached, the operation is terminated. The catalyst is regenerated by combustion in the usual manner and thereafter reused in the process including the start-up procedure as described. The deposition of coke and/or low product yield due to low catalyst activity usually are the factors which require termination of the operation. A space velocity (liquid volumes of feed, which include fresh feed and unconverted recycle feed in recycle operation, per volume of catalyst per hour) of between about 0.5 and 5 and preferably between about 0.75 and 2 may be used for start-up and/or the onstream reaction. A hydrogen (i.e. actual hydrogen content) rate of between 4,000 and 25,000 s.c.f./bbl. of feed and preferably between 7,000 and 18,000 s.c.f./bbl. of feed may be employed in the hydrocracking process.

At equilibrium, in recycle operation, some of the halogen and sulfur which are added and some of the ammonia derived from the feed are dissolved in the product. Also these materials are partly removed via the hydrogen bleed stream. Therefore they usually do not have to be scrubbed out of the hydrogen rich recycle stream unless the ammonia content of this stream becomes excessively high. The advantage of not scrubbing is of course a reduction in the amount of halogen and sulfur that needs to be added to the liquid hydrocarbon feed by the amount present in the hydrogen recycle.

EXAMPLE I This example shows the advantage of starting up with the catalyst initially in the oxide state as compared to starting with a presulfided catalyst and maintaining it in the sulfided condition by sulfur addition. The catalyst used was a pelleted 6 percent nickel-19 percent tungsten2 percent fluorine on Triple A silica-alumina catalyst. The feed was a hydrogen refined FCC furnace oil. In the first operation the start-up was with unpresulfided catalyst at below 450 F. The feed had the properties given in Table I, Column A. It was fortified to 2600 p.p.m. sulfur added as dimethyl disulfide and was charged along with 10,000 s.c.f. of 75 percent hydrogen/bbl. over the catalyst at 1.0 LHSV, and 1500 p.s.i.g. There was immediate recycle of unscrubbed hydrogen gas. However, there was no initial recycle of unconverted furnace oil. This amount of sulfur addition was theoretically sufficient to bring the catalyst to its lined-out maximum obtaintable sulfur content of about 5.5 percent (coke-free basis) in about 21 hours. Also at the end of the second day it was found that counting hydrogen sulfide returned in the recycle gas stream, the total sulfur charged to the reactor was about 1.13 percent based on the single-pass fresh feed. The temperature was increased to obtain 70 percent conversion of the charge to gasoline and at the end of two days this conversion was accomplished at 600 to 605 F. Recycle of unconverted feed to extinction was started after the second day. After 83 days of operation the temperature had been increased to 648 F. to maintain the desired 70 percent conversion of fresh feed plus recycle and the carbon on the catalyst was 2.72 percent. During the operation with recycle of unconverted feed, counting the hydrogen sulfide returned in the recycle gas, the total sulfur charged to the reactor 6 was about 0.79 percent based on total liquid reactor feed. In the comparison run the same catalyst was presulfided to the lined-out sulfur content of about 5.5 percent by subjecting it to an GHSV flow of an 82 percent hydrogen-18 percent H S mixture at 580 F. and 1200 p.s.i.g. for three hours. This catalyst was then used to process a hydrogen refined FCC furnace oil having the properties given in Table I, Column B, together with 1.38 percent sulfur added as tertiary butyl mercaptan and 10,000 s.c.f. of percent hyrogen/bbl. at 1000 p.s.i.g. and 1 LHSV in a single-pass operation. This amount of added sulfur made the reactor feed sulfur level about the same in this single-pass operation as it was at the end of two days in the preceding operation. Starting up was at 500 F. and no hot spot trouble was noted even in spite of the use of this relatively high temperature. At the end of two days the conversion of feed stock to gasoline at a temperature of 600 to 605 F. was only about 57 percent. The catalyst contained 13.39 percent carbon (after 4% days on stream).

EXAMPLE II In this example the advantage of our invention is shown by two runs in which the process of our invention is compared with a similar procedure using a substantially unsulfided catalyst. In one of these runs hydrogen pretreated light straight-nln furnace oil having the properties given in Table 1, Column C(a) was treated for 10 days. Thereafter the feed was changed to that shown in Column C(b) of Table I. The graphs shown in FIGURES 1, 2 and 3 marked Sulfur Addition were obtained as a re sult of operation in accordance with our invention. The start-up was at a temperature above the dew point of the charge but below 425 F. The hydrogen refined feed sulfided to 1662 p.p.m. sulfur with dimethyl disulfide and unconverted feed recycle after about three days was charged over the pelleted catalyst initially in the oxide form together with 18,000 s.c.f. of 75 percent hydrogen/bbl. at 0.94 LHSV and 750 p.s.i.g. and the temperature was gradually increased to obtain the desired con version of about 67 percent. Recycle of gas was started immediately and this gas was scrubbed to remove ammonia except for a short time during which it was established that no permanent poisoning of the catalyst resulted from not scrubbing out the ammonia. Hydrogen sulfide was not removed from the recycle gas. The amount of sulfur in the feed was theoretically adequate to cause catalyst sulfiding to the lined-out level of about 5.5 percent (coke free basis) in about 40 hours. After four days the amount of sulfur in fresh feed was cut to about 54 p.p.m. and after 19 days the source of added sulfur was changed from dimethyl disulfide to carbon disulfide which has been found to be the equivalent of dimethyl disulfide in this application. At equilibrium, the 54 p.p.m. sulfur, added to fresh feed and that returning as hydrogen sulfide in the recycle gas stream gave a total liquid reactor feed sulfur level of about 270 p.p.m. The graphs in FIGURES 1, 2 and 3 marked No Sulfur Addition were obtained in the same manner including the change in feed after 10 days, except that no extraneous sulfur was added to the pretreated feeds described in Table I, Columns C(a) and C(b), and the recycle gas was chemically scrubbed to remove both ammonia and any hydrogen sulfide. This 'amount of sulfur would result in such slow catalyst sulfiding that it would take several hundred days to reach the maximum lined-out catalyst sulfur content assuming that all the sulfur reacted, which is unlikely. In both operations 10 p.p.m. fluorine was added as ortho-fiuorotoluene to the fresh feed for the first 19 days of operation and thereafter fluorine addition was cut to 4 p.p.m. for catalyst activity and fluorine content maintenance. These amounts of fluorine amounted to 6.7 and 2.7 p.p.m. respectively based on total liquid reactor feed during recycle operation. It can be noticed that the operation with sulfur addition resulted in higher activity for the catalyst and a lower catalyst aging rate. Thus, not only is our proposed method more advantageous than starting up with presulfided catalyst, it is also better than starting up with unsulfided catalyst and not sulfiding it at as rapid a rate as proposed or only at the rate possible when processing normally hydrogen refined feeds. The catalysts employed in this example had the same nickel and tungsten content and were deposited on the same carrier as that described in Example 1.

EXAMPLE III In this example the feed stock described in Example I in connection with the unpresulfided catalyst was treated in the same manner as described in Example I in con nection with the unpresulfided catalyst, excepting that no sulfur was added to the feed stock until the operation had been continued for about 12 days. Thereafter sulfur addition to the fresh feed was started, and the reactor sulfur level counting hydrogen sulfide returned in the hydrogen rich recycle quickly lined-out at about 0.70 percent based on total liquid reactor feed. Prior to the sulfur addition, the aging rate of the catalyst was such as to require a temperature increase of 1.2 F. per day in order to maintain conversion. After the addition of sulfur was started, the daily temperature increase in order to maintain conversion amounted only to- 0.4 or 0.5 F. While this mode of operation results in poor activity and aging when processing feeds which are low in sulfur, nevertheless it is evident that by utilizing our invention including a low temperature start-up during the primary portion of the onstream reaction, substantially better results are obtained than if no sulfur is added. The catalyst employed in this example was the same as in Example I insofar as nickel and tungsten content and the carrier are concerned.

EXAMPLE IV In this example, start-up was at 450 F. with pelleted unpresulfided 6 percent nickel-19 percent tungsten-2 percent fluorine Triple A silica-alumina supported catalyst. Operation for the first three days was with the sulfur-free severely hydrogen refined FCC furnace oil described in Table I, Column D, and operation for the next day was with the essentially sulfur-free hydrogen refined FCC furnace oil described in Table I, Column B. During this period the temperature was increased to and maintained at 500 F. with other processing conditions being 1.0 LHSV, 1000 p.s.i.g. and 10,000 s.c.f. of 100 percent hydrogen/bbl. in single-pass operation. Subsequent to this sulfur-free break-in period during which the catalyst became partially reduced, 1.38 percent sulfur at t-butyl mercaptan was added to the essentially sulfur-free hydrogen refined FCC furnace oil described in Table I, Column B. After 4% days of processing this sulfur containing stock at 1.0 LHSV, 1000 p.s.i.g. and 10,0000 s.c.f. of 100 percent hydrogen/bbl. in singlepas operation the temperature had been lined-out at 600 to 605 F. and the conversion of feed stock to gasoline was about 74 percent. Also at this time, the catalyst attained its maximum lined-out sulfur content of about 5.5 percent (coke-free basis) and its carbon content was about 2.0 percent. This may be compared with the run in Example I, with the catalyst presulfided to the maximum lined-out sulfur content of about 5.5 percent and maintained at that sulfur content by processing the same feed fortified to 1.38 percent sulfur content by t-butyl mercaptan addition, where the conversion of stock to gasoline after 2 days was only 57 percent and after 4% days was only 51 percent. Furthermore, the catalyst contained 13.39 percent carbon at the end of the 4% days. Thus an advantageous modification of our operation as compared with the prior art can be to operate for a period of time under sulfur-free conditions during which time the catalyst is subjected to reducing conditions and then to cause sulfiding of the catalyst by addi- 8 tion of sulfur to the feed and prior to expiration of the major portion of the conversion period.

TABLE I C A B D Gravity, API 28. 4 28. 8 42. 4 41. 5 31. 2 Nitrogen, p.p.m l. 0 0. 2 4. 7 2. 2 0. 2 Sulfur, p.p.rn 5 1 17 3 1 Hydrocarbon type, v01.

percent:

406 416 375 389 412 612 630 537 549 638 440 448 407 415 448 490 490 432 436 492 percent 570 570 472 480 577 We claim:

1.. A start-up procedure for use in a hydrocracking process, which process comprises hydrocracking a hydrocarbon feed which is substantially composed of hydrocarbons having a boiling point below about 800 R, which has a nitrogen content below about 25 ppm. and which is substantially free of asphaltic materials utilizing a catalyst comprising essentially nickel-tungsten sulfide composited with a siliceous carrier having a high activity index, the nickel-tungsten components being substantially in sulfided form during the major conversion portion of the hydrocracking process, said start-up procedure comprising initially contacting the hydrocarbon feed with hydnogen in the presence of a catalyst consisting of nickel-tungsten composited with a siliceous carrier while the nickel and tungsten components are substantially unsulfided, at a hydrogen partial pressure of between about 400 and 3000 p.s.i., at an initial temperature below about 500 F. and While adding extraneous sulfur to the feed in an amount in the range of about 0.1 to about 2.0 percent by weight based upon the total liquid hydrocarbon feed, increasing the temperature until the desired conversion is obtained, reducing the amount of extraneous sulfur being added to the feed after the nickeltungsten components have been substantially converted into the sulfided form to an amount in the range of about 40 p.p.m. to about 1.5 percent by Weight based upon the total liquid hydrocarbon feed, maintaining the conversion by increasing the temperature and continuing the introduction of sulfur.

2. The start-up procedure of claim 1 wherein the catalyst contains at least about 0.1 percent by weight of combined halogen.

3. The start-up procedure of claim 1 wherein up to 20 p.p.m. of extraneous halogen is added to the feed.

4. The start-up procedure of claim 2 wherein up to 20 ppm. of extraneous halogen is added to the feed.

5. The start-up procedure of claim 1 wherein the carrier has an activity index of at least 40.

6. The start-up procedure of claim 1 wherein the nickeltungsten components of the catalyst when initially contacted with the hydrocarbon feed are les than about 50 percent sulfided.

7. The start-up procedure of claim 1 wherein the carrier is a silica-alumina carrier having an activity index of at least 45.

8. A start-up procedure for use in a hydrocracking process, which process comprises hydrocracking a hydrocarbon feed which is substantially composed of hydrocarbons boiling between about 350 and 800 E, which has a nitrogen content below about 25 ppm. and which is substantially free of asphaltic materials utilizing a cata lyst comprising essentially a nickel-tungsten sulfide corn posited with a silica-alumina carrier having an activity index of at least 45, the nickel-tungsten components being substantially in sulfided form during the major conversion portion of the hydrocracking process, said startup procedure comprising initially contacting the hydrocarbon feed with hydrogen in the presence of a catalyst consisting of nickel-tungsten composited with a silicaalumina carrier while the nickel and tungsten components are substantially unsulfided, at a hydrogen partial pressure of between about 400 and 1500 p.s.i., at an initial temperature below about 500 F. and while adding extraneous sulfur to the feed in an amount in the range of about 0.1 to about 2.0 percent by weight based upon the total liquid hydrocarbon feed, increasing the temperature at a rate which avoids localized zones of excessive temperature with resultant localized deposition of a relatively high percentage of coke, until a conversion between about 55 and 80 percent is obtained, reducing the amount of extraneous sulfur being added to the feed after the nickel and tungsten components have been substantially converted into the sulfided form to an amount in the range of about 40 ppm. to about 1.5 percent by weight based upon the total liquid hydrocarbon feed, maintaining the conversion between about 55 and 80 percent by increasing the temperature and continuing the introduction of sulfur.

9. The start-up procedure of claim 8 wherein the catalyst contains at least about 0.1 percent by weight of combined halogen.

10. The start-up procedure of claim 8 wherein up to 20 ppm. of extraneous halogen is added to the feed.

11. The start-up procedure of claim 9 wherein up to 20 ppm. of extraneous halogen is added to the feed.

12. The start-up procedure of claim 8 wherein the carrier has a surface area of at least 500 square meters per gram.

13. The start-up procedure of claim 8 wherein the nickel-tungsten components of the catalyst when initially contacted with the hydrocarbon feed are substantially in the oxide form.

14. A start-up procedure for use in a hydrocracking process, which process comprises hydrocracking a hydrocarbon feed 'which is substantially composed of hydrocarbons boiling between about 350 and 550 R, which has a nitrogen content below about ppm. and which is substantially free of asphaltic materials utilizing a catalyst comprising essentially a nickel-tungsten sulfide composited with a silicaaalumina carrier having an activity index of at least and a surface area of at least 500 square meters per gram, the catalyst containing at least about 0.1 percent by weight combined halogen and the nickel-tungsten components being substantially in sulfided form during the major conversion portion of the hydrocracking process, said start-up procedure comprising initially contacting the hydrocarbon feed with hydrogen in the presence of a catalyst consisting of nickeltungsten composited with a silica-alumina carrier while the nickel and tungsten component are substantially unsulfided, at a hydrogen partial pressure of between about 400 and 600 p.s.i., at an initial temperature below about 500 F. and while adding extraneous sulfur to the feed in an amount in the range of about 0.1 to about 2.0 percent by weight based upon the total liquid hydrocarbon feed, increasing the temperature at a rate which avoids localized zones of excessive temperature with resultant localized deposition of a relatively high percentage of coke, until a conversion between about and 80 percent is obtained, reducing the amount of extraneous sulfur being added to the feed after the nickel and tungsten components have been substantially converted into the sulfided form to an amount in the range of about 40 to about 1000 ppm. by weight based upon the total liquid hydrocarbon feed, maintaining the conversion between about 55 and 80 percent by increasing the temperature and continuing the introduction of sulfur.

15. The start-up procedure of claim 14 wherein up to 10 ppm. of extraneous halogen is added to the feed.

16. The start-up procedure of claim 14 wherein the initial contacting temperature is below about 450 F.

References Cited by the Examiner UNITED STATES PATENTS 2,934,492 4/ Hemminger et a1. 2081 10 2,944,005 7/60 Scott 208-109 3,099,617 7/ 63 Tulleners 208-1 10 ALPHONSO D. SULLIVAN, Primary Examiner.

UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 3, 213, 012 October 19, 1965 Robert E Kline et a1 Column 7 line 52, for "at" read as line 56, for "10,0000" read 10,000

Signed and seal-ed this 10th day of May 1966 (SEAL) Attest:

ERNEST W. SWIDER Attesting Officer Commissioner of Patents EDWARD J. BRENNER 

1. A START-UP PROCEDURE FOR USE IN A HYDROCRACKING PROCESS, WHICH PROCESS COMPRISES HYDROCRACKING A HYDROCARBON FEED WHICH IS SUBSTANTIALLY COMPOSED OF HYDROCARBONS HAVING A BOILING POINT BELOW ABOUT 800*F., WHICH HAS A NITROGEN CONTENT BELOW ABOUT 25 P.P.M. AND WHICH IS SUBSTANTIALLY FREE OF ASPHALTIC MATERIALS UTILIZING A CATALYST COMPRISING ESSENTIALLY NICKEL-TUNGSTEN SULFIDE COMPOSITED WITH A SILICEOUS CARRIER HAVING A HIGH ACTIVITY INDEX, THE NICKEL-TUNGSTEN COMPONENTS BEING SUBSTANTIALLY IN SULFIDED FORM DURING THE MAJOR CONVERSION PORTION OF THE HYDROCRACKING PROCESS, SAID START-UP PROCEDURE COMPRISING INITIALLY CONTACTING THE HYDROCARBON FEED WITH HYDROGEN IN THE PRESENCE OF A CATALYST CONSISTING OF NICKEL-TUNGSTEN COMPOSITED WITH A SILICEOUS CARRIER WHILE THE NICKEL AND TUNGSTEN COMPONENTS ARE SUBSTANTIALLY UNSULFIDED, AT A HYDROGEN PARTIAL PRESSURE OF BETWEEN ABOUT 400 AND 3000 P.S.I., AT AN INITIAL TEMPERATURE BELOW ABOUT 500*F. AND WHILE ADDING EXTRANEOUS SULFUR TO THE FEED IN AN AMOUNT IN THE RANGE OF ABOUT 0.1 TO ABOUT 2.0 PERCENT BY WEIGHT BASED UPON THE TOTAL LIQUID HYDROCARBON FEED, INCREASING THE TEMPERATURE UNTIL THE DESIRED CONVERSION IS OBTAINED, REDUCING THE AMOUNT OF EXTRANEOUS SULFUR BEING ADDED TO THE FEED AFTER THE NICKELTUNGSTEN COMPONENTS HAVE BEEN SUBSTANTIALLY CONVERTED INTO THE SULFIDED FORM TO AN AMOUNT IN THE RANGE OF ABOUT 40 P.P.M. TO ABOUT 1.5 PERCENT BY WEIGHT BASED UPON THE TOTAL LIQUID HYDROCARBON FEED, MAINTAINING THE CONVERSION BY INCREASING THE TEMPERATURE AND CONTINUING THE INTRODUCTION OF SULFUR. 